1. Field of Endeavor
The present invention relates generally to process control. More particularly the present invention relates to techniques for enhanced control of processes, such as those utilized for air pollution control. Examples of such processes include but are not limited to wet and dry flue gas desulfurization (WFGD/DFGD), nitrogen oxide removal via selective catalytic reduction (SCR), and particulate removal via electrostatic precipitation (ESP).
2. Background
Wet Flue Gas Desulfurization:
As noted, there are several air pollution control processes, to form a basis for discussion; the WFGD process will be highlighted. The WFGD process is the most commonly used process for removal of SO2 from flue gas in the power industry. FIG. 1, is a block diagram depicting an overview of a wet flue gas desulfurization (WFGD) subsystem for removing SO2 from the dirty flue gas, such as that produced by fossil fuel, e.g. coal, fired power generation systems, and producing a commercial grade byproduct, such as one having attributes which will allow it to be disposed of at a minimized disposal cost, or one having attributes making it saleable for commercial use.
In the United States of America, the presently preferred byproduct of WFGD is commercial grade gypsum having a relatively high quality (95+% pure) suitable for use in wallboard, which is in turn used in home and office construction. Commercial grade gypsum of high quality (˜92%) is also the presently preferred byproduct of WFGD in the European Union and Asia, but is more typically produced for use in cement, and fertilizer. However, should there be a decline in the market for higher quality gypsum, the quality of the commercial grade gypsum produced as a byproduct of WFGD could be reduced to meet the less demanding quality specifications required for disposal of at minimum costs. In this regard, the cost of disposal may be minimized if, for example, the gypsum quality is suitable for either residential landfill or for backfilling areas from which the coal utilized in generating power has been harvested.
As shown in FIG. 1, dirty, SO2 laden flue gas 112 is exhausted from a boiler or economizer (not shown) of a coal fired power generation system 110 to the air pollution control system (APC) 120. Commonly the dirty flue gas 112 entering the APC 120 is not only laden with SO2, but also contains other so called pollutants such as NOx and particulate matter. Before being processed by the WFGD subsystem, the dirty flue gas 112 entering the APC 120 is first directed to other APC subsystems 122 in order remove NOx and particulate matter from the dirty flue gas 112. For example, the dirty flue gas may be processed via a selective catalytic reduction (SCR) subsystem (not shown) to remove NOx and via an electrostatic precipitator subsystem (EPS) (not shown) or filter (not shown) to remove particulate matter.
The SO2 laden flue gas 114 exhausted from the other APC subsystems 122 is directed to the WFGD subsystem 130. SO2 laden flue gas 114 is processed by the absorber tower 132. As will be understood by those skilled in the art, the SO2 in the flue gas 114 has a high acid concentration. Accordingly, the absorber tower 132 operates to place the SO2 laden flue gas 114 in contact with liquid slurry 148 having a higher pH level than that of the flue gas 114.
It will be recognized that most conventional WFGD subsystems include a WFGD processing unit of the type shown in FIG. 1. This is true, for many reasons. For example, as is well understood in the art, WFGD processing units having a spray absorber towers have certain desirable process characteristics for the WFGD process. However, WFGD processing units having other absorption/oxidation equipment configurations could, if desired, be utilized in lieu of that shown in FIG. 1 and still provide similar flue gas desulfurization functionality and achieve similar benefits from the advanced process control improvements presented in this application. For purposes of clarity and brevity, this discussion will reference the common spray tower depicted in FIG. 1, but it should be noted that the concepts presented could be applied to other WFGD configurations.
During processing in the countercurrent absorber tower 132, the SO2 in the flue gas 114 will react with the calcium carbonate-rich slurry (limestone and water) 148 to form calcium sulfite, which is basically a salt and thereby removing the SO2 from the flue gas 114. The SO2 cleaned flue gas 116 is exhausted from the absorber tower 132, either to an exhaust stack 117 or to down-steam processing equipment (not shown). The resulting transformed slurry 144 is directed to the crystallizer 134, where the salt is crystallized. The crystallizer 134 and the absorber 132 typically reside in a single tower with no physical separation between them—while there are different functions (absorption in the gas phase and crystallization in the liquid phase) going on, the two functions occur in the same process vessel. From here, gypsum slurry 146, which includes the crystallized salt, is directed from the crystallizer 134 to the dewatering unit 136. Additionally, recycle slurry 148, which may or may not include the same concentration of crystallized salts as the gypsum slurry 146, is directed from the crystallizer 134 through pumps 133 and back to the absorber tower 132 to continue absorption cycle.
The blower 150 pressurizes ambient air 152 to create oxidation air 154 for the crystallizer 134. The oxidation air 154 is mixed with the slurry in the crystallizer 134 to oxidize the calcium sulfite to calcium sulfate. Each molecule of calcium sulfate binds with two molecules of water to form a compound that is commonly referred to as gypsum 160. As shown, the gypsum 160 is removed from the WFGD processing unit 130 and sold to, for example manufacturers of construction grade wallboard.
Recovered water 167, from the dewatering unit 136 is directed to the mixer/pump 140 where it is combined with fresh ground limestone 174 from the grinder 170 to create limestone slurry. Since some process water is lost to both the gypsum 160 and the waste stream 169, additional fresh water 162, from a fresh water source 164, is added to maintain the limestone slurry density. Additionally, waste, such as ash, is removed from the WFGD processing unit 130 via waste stream 169. The waste could, for example, be directed to an ash pond or disposed of in another manner.
In summary, the SO2 within the SO2 laden flue gas 114 is absorbed by the slurry 148 in the slurry contacting area of the absorber tower 132, and then crystallized and oxidized in the crystallizer 134 and dewatered in the dewatering unit 136 to form the desired process byproduct, which in this example, is commercial grade gypsum 160. The SO2 laden flue gas 114 passes through the absorber tower 132 in a matter of seconds. The complete crystallization of the salt within the transformed slurry 144 by the crystallizer 134 may require from 8 hours to 20+ hours. Hence, the crystallizer 134 has a large volume that serves as a slurry reservoir crystallization. The recycle slurry 148 is pumped back to the top of the absorber to recover additional SO2.
As shown, the slurry 148 is fed to an upper portion of the absorber tower 132. The tower 132 typically incorporates multiple levels of spray nozzles to feed the slurry 148 into the tower 132. The absorber 132, is operated in a countercurrent configuration: the slurry spray flows downward in the absorber and comes into contact with the upward flowing SO2 laden flue gas 114 which has been fed to a lower portion of the absorber tower.
Fresh limestone 172, from limestone source 176, is first ground in the grinder 170 (typically a ball mill) and then mixed with (recovered water 167 and fresh/make-up water 162 in a mixer 140 to form limestone slurry 141. The flow of the ground limestone 174 and water 162 via valve (not shown) to the mixer/tank 140 are controlled to maintain a sufficient inventory of fresh limestone slurry 141 in the mixer/tank 140. The flow of fresh limestone slurry 141 to the crystallizer 134 is adjusted to maintain an appropriate pH for the slurry 148, which in turn controls the amount of SO2 removed from the flue gas 114. WFGD processing typically accomplishes 92–97% removal of SO2 from the flue gas, although those skilled in the art will recognize that but utilizing certain techniques and adding organic acids to the slurry the removal of SO2 can increase to greater than 97%.
As discussed above, conventional WFGD subsystems recycle the slurry. Although some waste water and other waste will typically be generated in the production of the gypsum, water is reclaimed to the extent possible and used to make up fresh limestone slurry, thereby minimizing waste and costs, which would be incurred to treat the process water.
It will be recognized that because limestone is readily available in large quantities in most locations, it is commonly used as the reactant in coal gas desulfurization processing. However, other reactants, such as quick lime or a sodium compound, could alternatively be used, in lieu of limestone. These other reactants are typically more expensive and are not currently cost-competitive with the limestone reactant. However, with very slight modifications to the mixer 140 and upstream reactant source, an existing limestone WFGD could be operated using quick lime or a sodium compound. In fact, most WFGD systems include a lime backup subsystem so the WFGD can be operated if there are problems with limestone delivery and/or extended maintenance issues with the grinder 170.
FIG. 2 further details certain aspects of the WFGD subsystem shown in FIG. 1. As shown, the dewatering unit 136 may include both a primary dewatering unit 136A and a secondary dewatering unit 136B. The primary dewatering unit 136A preferably includes hydrocyclones for separating the gypsum and water. The secondary dewatering unit 136B preferably includes a belt dryer for drying the gypsum. As has been previously discussed, the flue gas 114 enters the absorber 132, typically from the side, and flows upward through a limestone slurry mist that is sprayed into the upper portion of the absorber tower. Prior to exiting the absorber, the flue gas is put through a mist eliminator (ME) (not shown) that is located in the top of the absorber 132; the mist eliminator removes entrained liquid and solids from the flue gas stream. To keep the mist eliminator clean of solids, a ME water wash 200 applied to the mist eliminator. As will be understood, the ME wash 200 keeps the ME clean within the absorber tower 132 with water from the fresh water source 164. The ME wash water 200 is the purest water fed to the WFGD subsystem 130.
As noted above, the limestone slurry mist absorbs a large percentage of the SO2 (e.g., 92–97%) from the flue gas that is flowing through the absorber tower 132. After absorbing the SO2, the slurry spray drops to the crystallizer 134. In a practical implementation, the absorber tower 132 and the crystallizer 134 are often housed in a single unitary structure, with the absorber tower located directly above the crystallizer within the structure. In such implementations, the slurry spray simply drops to the bottom of the unitary structure to be crystallized.
The limestone slurry reacts with the SO2 to produce gypsum (calcium sulfate dehydrate) in the crystallizer 134. As previously noted, forced, compressed oxidation air 154 is used to aid in oxidation, which occurs in the following reaction:SO2+CaCO3+½O2+2H2O—>CaSO4.2H2O+CO2  (1)The oxidation air 154 is forced into the crystallizer 134, by blower 150. Oxidation air provides additional oxygen needed for the conversion of the calcium sulfite to calcium sulfate.
The absorber tower 132 is used to accomplish the intimate flue gas/liquid slurry contact necessary to achieve the high removal efficiencies required by environmental specifications. Countercurrent open-spray absorber towers provide particularly desirable characteristics for limestone-gypsum WFGD processing: they are inherently reliable, have lower plugging potential than other tower-based WFGD processing unit components, induce low pressure drop, and are cost-effective from both a capital and an operating cost perspective.
As shown in FIG. 2, the water source 164 typically includes a water tank 164A for storing a sufficient quantity of fresh water. Also typically included there is one or more pumps 164B for pressurizing the ME wash 200 to the absorber tower 132, and one or more pumps 164C for pressurizing the fresh water flow 162 to the mixer 140. The mixer 140 includes a mixing tank 140A and one more slurry pumps 140B to move the fresh limestone slurry 141 to the crystallizer 134. One or more additional very large slurry pumps 133 (see FIG. 1) are required to lift the slurry 148 from the crystallizer 134 to the multiple spray levels in the top of the absorber tower 132.
As will be described further below, typically, the limestone slurry 148 enters the absorber tower 132, via spray nozzles (not shown) disposed at various levels of the absorber tower 132. When at full load, most WFGD subsystems operate with at least one spare slurry pump 133. At reduced loads, it is often possible to achieve the required SO2 removal efficiency with a reduced number of slurry pumps 133. There is significant economic incentive to reduce the pumping load of the slurry pumps 133. These pumps are some of the largest pumps in the world and they are driven by electricity that could otherwise be sold directly to the power grid (parasitic power load).
The gypsum 160 is separated from liquids in the gypsum slurry 146 in the primary dewaterer unit 136A, typically using a hydrocyclone. The overflow of the hydrocyclone, and/or one or more other components of primary dewaterer unit 136A, contains a small amount of solids. As shown in FIG. 2, this overflow slurry 146A is returned to the crystallizer 134. The recovered water 167 is sent back to mixer 140 to make fresh limestone slurry. The other waste 168 is commonly directed from the primary dewaterer unit 136A to an ash pond 210. The underflow slurry 202 is directed to the secondary dewaterer unit 136B, which often takes the form of a belt filter, where it is dried to produce the gypsum byproduct 160. Again, recovered water 167 from the secondary dewaterer unit 136B is returned to the mixer/pump 140. As shown in FIG. 1, hand or other gypsum samples 161 are taken and analyzed, typically every few hours, to determine the purity of the gypsum 160. No direct on-line measurement of gypsum purity is conventionally available.
As shown in FIG. 1, a proportional integral derivative (PID) controller 180 is conventionally utilized in conjunction with a feedforward controller (FF) 190 to control the operation of the WFDG subsystem. Historically, PID controllers directed pneumatic analog control functions. Today, PID controllers direct digital control functions, using mathematically formulations. The goal of FF 190/PID controller 180 is to control the slurry pH, based on an established linkage. For example, there could be an established linkage between the adjustment of valve 199 shown in FIG. 1, and a measured pH value of slurry 148 flowing from the crystallizer 134 to the absorber tower 132. If so, valve 199 is controlled so that the pH of the slurry 148 corresponds to a desired value 186, often referred to as a setpoint (SP).
The FF 190/PID controller 180 will adjust the flow of the limestone slurry 141 through valve 199, based on the pH setpoint, to increase or decrease the pH value of the slurry 148 measured by the pH sensor 182. As will be understood, this is accomplish by the FF/PID controller transmitting respective control signals 181 and 191, which result in a valve adjustment instruction, shown as flow control SP 196, to a flow controller which preferably is part of the valve 199. Responsive to flow control SP 196, the flow controller in turn directs an adjustment of the valve 199 to modify the flow of the limestone slurry 141 from the mixer/pump 140 to the crystallizer 134.
The present example shows pH control using the combination of the FF controller 190 and the PID controller 180. Some installations will not include the FF controller 190.
In the present example, the PID controller 180 generates the PID control signal 181 by processing the measured slurry pH value 183 received from the pH sensor 182, in accordance with a limestone flow control algorithm representing an established linkage between the measured pH value 183 of the slurry 148 flowing from the crystallizer 134 to the absorber tower 132. The algorithm is typically stored at the PID controller 180, although this is not mandatory. The control signal 181 may represent, for example, a valve setpoint (VSP) for the valve 199 or for a measured value setpoint (MVSP) for the flow of the ground limestone slurry 141 exiting the valve 199.
As is well understood in the art, the algorithm used by the PID controller 180 has a proportional element, an integral element, and a derivative element. The PID controller 180 first calculates the difference between the desired SP and the measured value, to determine an error. The PID controller next applies the error to the proportional element of the algorithm, which is an adjustable constant for the PID controller, or for each of the PID controllers if multiple PID controllers are used in the WFGD subsystem. The PID controller typically multiples a tuning factor or process gain by the error to obtain a proportional function for adjustment of the valve 199.
However, if the PID controller 180 does not have the correct value for the tuning factor or process gain, or if the process conditions are changing, the proportional function will be imprecise. Because of this imprecision, the VSP or MVSP generated by the PID controller 180 will actually have an offset from that corresponding to the desired SP. Accordingly, the PID controller 180 applies the accumulated error over time using the integral element. The integral element is a time factor. Here again, the PID controller 180 multiplies a tuning factor or process gain by the accumulated error to eliminate the offset.
Turning now to the derivative element. The derivative element is an acceleration factor, associated with continuing change. In practice, the derivative element is rarely applied in PID controllers used for controlling WFGD processes. This is because application of the derivative element is not particularly beneficial for this type of control application. Thus, most controllers used for in WFGD subsystems are actually PI controllers. However, those skilled in the art will recognize that, if desired, the PID controller 180 could be easily configured with the necessary logic to apply a derivative element in a conventional manner.
In summary, there are three tuning constants, which may be applied by conventional PID controllers to control a process value, such as the pH of the recycle slurry 148 entering the absorber tower 132, to a setpoint, such as the flow of fresh lime stone slurry 141 to the crystallizer 134. Whatever setpoint is utilized, it is always established in terms of the process value, not in terms of a desired result, such as a value of SO2 remaining in the flue gas 116 exhausted from the absorber tower 132. Stated another way, the setpoint is identified in process terms, and it is necessary that the controlled process value be directly measurable in order for the PID controller to be able to control it. While the exact form of the algorithm may change from one equipment vendor to another, the basic PID control algorithm has been in use in the process industries for well over 75 years.
Referring again to FIGS. 1 and 2, based on the received instruction from the PID controller 180 and the FF controller 190, the flow controller generates a signal, which causes the valve 199 to open or close, thereby increasing or decreasing the flow of the ground limestone slurry 141. The flow controller continues control of the valve adjustment until the valve 199 has been opened or closed to match the VSP or the measured value of the amount of limestone slurry 141 flowing to from the valve 1992 matches the MVSP.
In the exemplary conventional WFGD control described above, the pH of the slurry 148 is controlled based on a desired pH setpoint 186. To perform the control, the PID 180 receives a process value, i.e. the measured value of the pH 183 of the slurry 148, from the sensor 182. The PID controller 180 processes the process value to generate instructions 181 to the valve 199 to adjust the flow of fresh limestone slurry 141, which has a higher pH than the crystallizer slurry 144, from the mixer/tank 140, and thereby adjust the pH of the slurry 148. If the instructions 181 result in a further opening of the valve 199, more limestone slurry 141 will flow from the mixer 140 and into the crystallizer 134, resulting in an increase in the pH of the slurry 148. On the other hand, if the instructions 181 result in a closing of the valve 199, less limestone slurry 141 will flow from the mixer 140 and therefore into the crystallizer 134, resulting in a decrease in the pH of the slurry 148.
Additionally, the WFGD subsystem may incorporate a feed forward loop, which is implemented using a feed forward unit 190 in order to ensure stable operation. As shown in FIG. 1, the concentration value of SO2 189 in the flue gas 114 entering the absorber tower 132 is measured by sensor 188 and input to the feed forward unit 190. Many WFGD systems that include the FF control element may combine the incoming flue gas SO2 concentration 189 with a measure of generator load from the Power Generation System 110, to determine the quantity of inlet SO2 rather than just the concentration and, then use this quantity of inlet SO2 as the input to FF 190. The feed forward unit 190 serves as a proportional element with a time delay.
In the exemplary implementation under discussion, the feed forward unit 190 receives a sequence of SO2 measurements 189 from the sensor 188. The feed forward unit 190 compares the currently received concentration value with the concentration value received immediately preceding the currently received value. If the feed forward unit 190 determines that a change in the measured concentrations of SO2 has occurred, for example from 1000–1200 parts per million, it is configured with the logic to smooth the step function, thereby avoiding an abrupt change in operations.
The feed forward loop dramatically improves the stability of normal operations because the relationship between the pH value of the slurry 148 and the amount of limestone slurry 141 flowing to the crystallizer 134 is highly nonlinear, and the PID controller 180 is effectively a linear controller. Thus, without the feed forward loop, it is very difficult for the PID 180 to provide adequate control over a wide range of pH with the same tuning constants.
By controlling the pH of the slurry 148, the PID controller 180 effects both the removal of SO2 from the SO2 laden flue gas 114 and the quality of the gypsum byproduct 160 produced by the WFGD subsystem. Increasing the slurry pH by increasing the flow of fresh limestone slurry 141 increases the amount of SO2 removed from the SO2 laden flue gas 114. On the other hand, increasing the flow of limestone slurry 141, and thus the pH of the slurry 148, slows the SO2 oxidation after absorption, and thus the transformation of the calcium sulfite to sulfate, which in turn will result in a lower quality of gypsum 160 being produced.
Thus, there are conflicting control objectives of removing SO2 from the SO2 laden flue gas 114, and maintaining the required quality of the gypsum byproduct 160. That is, there may be a conflict between meeting the SO2 emission requirements and the gypsum quality requirements.
FIG. 3 details further aspects of the WFGD subsystem described with reference to FIGS. 1 and 2. As shown, SO2 laden flue gas 114 enters into a bottom portion of the absorber tower 132 via an aperture 310, and SO2 free flue gas 116 exits from an upper portion of the absorber tower 132 via an aperture 312. In this exemplary conventional implementation, a counter current absorber tower is shown, with multiple slurry spray levels. As shown, the ME wash 200 enters the absorber tower 132 and is dispersed by wash sprayers (not shown).
Also shown are multiple absorber tower slurry nozzles 306A, 306B and 306C, each having a slurry sprayer 308A, 308B or 308C, which sprays slurry into the flue gas to absorb the SO2. The slurry 148 is pumped from the crystallizer 134 shown in FIG. 1, by multiple pumps 133A, 133B and 133C, each of which pumps the slurry up to a different one of the levels of slurry nozzles 306A, 306B or 306C. It should be understood that although 3 different levels of slurry nozzles and sprayers are shown, the number of nozzles and sprayers would vary depending on the particular implementation.
A ratio of the flow rate of the liquid slurry 148 entering the absorber 132 over the flow rate of the flue gas 116 leaving the absorber 132 is commonly characterized as the L/G. L/G is one of the key design parameters in WFGD subsystems.
The flow rate of the flue gas 116 (saturated with vapor), designated as G, is a function of inlet flue gas 112 from the power generation system 110 upstream of the WFGD processing unit 130. Thus, G is not, and cannot be, controlled, but must be addressed, in the WFGD processing. So, to impact L/G, the “L” must be adjusted. Adjusting the number of slurry pumps in operation and the “line-up” of these slurry pumps controls the flow rate of the liquid slurry 148 to the WFGD absorber tower 132, designated as L. For example, if only two pumps will be run, running the pumps to the upper two sprayer levels vs. the pumps to top and bottom sprayer levels will create different “L”s.
It is possible to adjust “L” by controlling the operation of the slurry pumps 133A, 133B and 133C. Individual pumps may be turned on or off to adjust the flow rate of the liquid slurry 148 to the absorber tower 132 and the effective height at which the liquid slurry 148 is introduced to the absorber tower. The higher the slurry is introduced into the tower, the more contact time it has with the flue gas resulting in more SO2 removal, but this additional SO2 removal comes at the penalty of increased power consumption to pump the slurry to the higher spray level. It will be recognized that the greater the number of pumps, the greater the granularity of such control.
Pumps 133A–133C, which are extremely large pieces of rotating equipment, can be started and stopped automatically or manually. Most often, in the USA, these pumps are controlled manually by the subsystem operator. It is more common to automate starting/stopping rotating equipment, such as pumps 133A–133C in Europe.
If the flow rate of the flue gas 114 entering the WFGD processing unit 130 is modified due to a change in the operation of the power generation system 110, the WFGD subsystem operator may adjust the operation of one or more of the pumps 133A–133C. For example, if the flue gas flow rate were to fall to 50% of the design load, the operator, or special logic in the control system, might shut down one or more of the pumps that pump slurry to the spray level nozzles at one or more spray level.
Although not shown in FIG. 3, it will be recognized that extra spray levels, with associated pumps and slurry nozzles, are often provided for use during maintenance of another pump, or other slurry nozzles and/or slurry sprayers associated with the primary spray levels. The addition of this extra spray level adds to the capital costs of the absorber tower and hence the subsystem. Accordingly, some WFGD owners will decide to eliminate the extra spray level and to avoid this added capital costs, and instead add organic acids to the slurry to enhance its ability to absorb and therefore remove SO2 from the flue gas during such maintenance periods. However, these additives tend to be expensive and therefore their use will result in increased operational costs, which may, over time, offset the savings in capital costs.
As indicated in Equation 1 above, to absorb SO2, a chemical reaction must occur between the SO2 in the flue gas and the limestone in the slurry. The result of the chemical reaction in the absorber is the formation of calcium sulfite. In the crystallizer 134, the calcium sulfite is oxidized to form calcium sulfate (gypsum). During this chemical reaction, oxygen is consumed. To provide sufficient oxygen and enhance the speed of the reaction, additional O2 is added by blowing compressed air 154 into the liquid slurry in the crystallizer 134.
More particularly, as shown in FIG. 1 ambient air 152 is compressed to form compressed air 154, and forced into the crystallizer 134 by a blower, e.g. fan, 150 in order to oxidize the calcium sulfite in the recycle slurry 148 which is returned from the crystallizer 134 to the absorber 132 and the gypsum slurry 146 sent to the dewatering system 136 for further processing. To facilitate adjustment of the flow of oxidation air 154, the blower 150 may have a speed or load control mechanism.
Preferably, the slurry in the crystallizer 134 has excess oxygen. However, there is an upper limit to the amount of oxygen that can be absorbed or held by slurry. If the O2 level within the slurry becomes too low, the chemical oxidation of CaSO3 to CaSO4 in the slurry will cease. When this occurs, it is commonly referred to as limestone blinding. Once limestone blinding occurs, limestone stops dissolving into the slurry solution and SO2 removal can be dramatically reduced. The presence of trace amounts of some minerals can also dramatically slow the oxidation of calcium sulfite and/or limestone dissolution to create limestone blinding.
Because the amount of O2 that is dissolved in the slurry is not a measurable parameter, slurry can become starved for O2 in conventional WFGD subsystems if proper precautions are not taken. This is especially true during the summer months when the higher ambient air temperature lowers the density of the ambient air 152 and reduces the amount of oxidation air 154 that can be forced into the crystallizer 134 by the blower 150 at maximum speed or load. Additionally, if the amount of SO2 removed from the flue gas flow increases significantly, a corresponding amount of additional O2 is required to oxidize the SO2. Thus, the slurry can effectively become starved for O2 because of an increase in the flow of SO2 to the WFGD processing unit.
It is necessary to inject compressed air 154 that is sufficient, within design ratios, to oxidize the absorbed SO2. If it is possible to adjust blower 150 speed or load, and turning down the blower 150 at lower SO2 loads and/or during cooler ambient air temperature periods is desirable because it saves energy. When the blower 150 reaches maximum load, or all the O2 of a non-adjustable blower 150 is being utilized, it is not possible to oxidize an incremental increase in SO2. At peak load, or without a blower 150 speed control that accurately tracks SO2 removal, it is possible to create an O2 shortage in the crystallizer 134.
However, because it is not possible to measure the O2 in the slurry, the level of O2 in the slurry is not used as a constraint on conventional WFGD subsystem operations. Thus, there is no way of accurately monitoring when the slurry within the crystallizer 134 is becoming starved for O2. Accordingly, operators, at best, will assume that the slurry is becoming starved for O2 if there is a noticeable decrease in the quality of the gypsum by-product 160, and use their best judgment to control the speed or load of blower 150 and/or decrease SO2 absorption efficiency to balance the O2 being forced into the slurry, with the absorbed SO2 that must be oxidized. Hence, in conventional WFGD subsystems balancing of the O2 being forced into the slurry with the SO2 required to be absorbed from the flue gas is based, at best, on operator judgment.
In summary, conventional control of large WFGD subsystems for utility application is normally carried out within a distributed control system (DCS) and generally consists of on-off control logic as well as FF/PID feedback control loops. The parameters controlled are limited to the slurry pH level, the L/G ratio and the flow of forced oxidation air.
The pH must be kept within a certain range to ensure high solubility of SO2 (i.e. SO2 removal efficiency) high quality (purity) gypsum, and prevention of scale buildup. The operating pH range is a function of equipment and operating conditions. The pH is controlled by adjusting the flow of fresh limestone slurry 141 to the crystallizer 134. The limestone slurry flow adjustment is based on the measured pH of the slurry detected by a sensor. In a typically implementation, a PID controller and, optionally, FF controller included in the DCS are cascaded to a limestone slurry flow controller. The standard/default PID algorithm is used for pH control application.
The liquid-to-gas ratio (L/G) is the ratio of the liquid slurry 148 flowing to the absorber tower 132 to the flue gas flow 114. For a given set of subsystem variables, a minimum L/G ratio is required to achieve the desired SO2 absorption, based on the solubility of SO2 in the liquid slurry 148. The L/G ratio changes either when the flue gas 114 flow changes, or when the liquid slurry 148 flow changes, which typically occurs when slurry pumps 133 are turned on or off.
The oxidation of calcium sulfite to form calcium sulfate, i.e. gypsum, is enhanced by forced oxidation, with additional oxygen in the reaction tank of the crystallizer 134. Additional oxygen is introduced by blowing air into the slurry solution in the crystallizer 134. With insufficient oxidation, sulfite—limestone blinding can occur resulting in poor gypsum quality, and potentially subsequent lower SO2 removal efficiency, and a high chemical oxygen demand (COD) in the waste water.
The conventional WFGD process control scheme is comprised of standard control blocks with independent rather than integrated objectives. Currently, the operator, in consultation with the engineering staff, must try to provide overall optimal control of the process. To provide such control, the operator must take the various goals and constraints into account.
Minimized WFGD Operation Costs—Power plants are operated for no other reason than to generate profits for their owners. Thus, it is beneficial to operate the WFGD subsystem at the lowest appropriate cost, while respecting the process, regulatory and byproduct quality constraints and the business environment.
Maximize SO2 Removal Efficiency—Clean air regulations establish SO2 removal requirements. WFGD subsystems should be operated to remove SO2 as efficiently as appropriate, in view of the process, regulatory and byproduct quality constraints and the business environment.
Meet Gypsum Quality Specification—The sale of gypsum as a byproduct mitigates WFGD operating costs and depends heavily on the byproduct purity meeting a desired specification. WFGD subsystems should be operated to produce a gypsum byproduct of an appropriate quality, in view of the process, regulatory and byproduct quality constraints and the business environment.
Prevent Limestone Blinding—Load fluctuations and variations in fuel sulfur content can cause excursions in SO2 in the flue gas 114. Without proper compensating adjustments, this can lead to high sulfite concentrations in the slurry, which in turn results in limestone blinding, lower absorber tower 132 SO2 removal efficiency, poor gypsum quality, and a high chemical oxygen demand (COD) in the wastewater. WFGD subsystems should be operated to prevent limestone binding, in view of the process constraints.
In a typical operational sequence, the WFGD subsystem operator determines setpoints for the WFGD process to balance these competing goals and constraints, based upon conventional operating procedures and knowledge of the WFGD process. The setpoints commonly include pH, and the operational state of the slurry pumps 133 and oxidation air blower 150.
There are complex interactions and dynamics in the WFGD process; as a result, the operator selects conservative operating parameters so that the WFGD subsystem is able to meet/exceed hard constraints on SO2 removal and gypsum purity. In making these conservative selections, the operator often, if not always, sacrifices minimum-cost operation.
For example, FIG. 4 shows SO2 removal efficiency and gypsum purity as a function of pH. As pH is increased, the SO2 removal efficiency increases, however, the gypsum purity decreases. Since the operator is interested in improving both SO2 removal efficiency and gypsum purity, the operator must determine a setpoint for the pH that is a compromise between these competing goals.
In addition, in most cases, the operator is required to meet a guaranteed gypsum purity level, such as 95% purity. Because of the complexity of the relationships shown in FIG. 4, the lack of direct on-line measurement of gypsum purity, the long time dynamics of gypsum crystallization, and random variations in operations, the operator often chooses to enter a setpoint for pH that will guarantee that the gypsum purity level is higher than the specified constraint under any circumstances. However, by guaranteeing the gypsum purity, the operator often sacrifices the SO2 removal efficiency. For instance, based upon the graph in FIG. 4, the operator may select a pH of 5.4 to guarantee of 1% cushion above the gypsum purity constraint of 95%. However, by selecting this setpoint for pH, the operator sacrifices 3% of the SO2 removal efficiency.
The operator faces similar compromises when SO2 load, i.e. the flue gas 114 flow, drops from full to medium. At some point during this transition, it may be beneficial to shut off one or more slurry pumps 133 to save energy, since continued operation of the pump may provide only slightly better SO2 removal efficiency. However, because the relationship between the power costs and SO2 removal efficiency is not well understood by most operators, operators will typically take a conservative approach. Using such an approach, the operators might not adjust the slurry pump 133 line-up, even though it would be more beneficial to turn one or more of the slurry pumps 133 off.
It is also well known that many regulatory emission permits provide for both instantaneous emission limits and some form of rolling-average emission limits. The rolling-average emission limit is an average of the instantaneous emissions value over some moving, or rolling, time-window. The time-window may be as short as 1-hour or as long as 1-year. Some typical time-windows are 1-hour, 3-hours, 8-hours, 24-hours, 1-month, and 1-year. To allow for dynamic process excursions, the instantaneous emission limit is typically higher than rolling average limit. However, continuous operation at the instantaneous emission limit will result in a violation of the rolling-average limit.
Conventionally, the PID 180 controls emissions to the instantaneous limit, which is relatively simple. To do this, the operating constraint for the process, i.e. the instantaneous value, is set well within the actual regulatory emission limit, thereby providing a safety margin.
On the other hand, controlling emissions to the rolling-average limit is more complex. The time-window for the rolling-average is continually moving forward. Therefore, at any given time, several time-windows are active, spanning one time window from the given time back over a period of time, and another time window spanning from the given time forward over a period of time.
Conventionally, the operator attempts to control emissions to the rolling-average limit, by either simply maintaining a sufficient margin between the operating constraint set in the PID 180 for the instantaneous limit and the actual regulatory emission limit, or by using operator judgment to set the operating constraint in view of the rolling-average limit. In either case, there is no explicit control of the rolling-average emissions, and therefore no way to ensure compliance with the rolling-average limit or prevent costly over-compliance.
Selective Catalytic Reduction System:
Briefly turning to another exemplary air pollution control process, the selective catalytic reduction (SCR) system for NOx removal, similar operating challenges can be identified. An overview of the SCR process is shown in FIG. 20.
The following process overview is from “Control of Nitrogen Oxide Emissions: Selective Catalytic Reduction (SCR)”, Topical Report Number 9, Clean Coal Technology, U.S Dept. of Energy, 1997:
Process Overview
NOx, which consists primarily of NO with lesser amounts of NO2, is converted to nitrogen by reaction with NH3 over a catalyst in the presence of oxygen. A small fraction of the SO2, produced in the boiler by oxidation of sulfur in the coal, is oxidized to sulfur trioxide (SO3) over the SCR catalyst. In addition, side reactions may produce undesirable by-products: ammonium sulfate, (NH4)2SO4, and ammonium bisulfate, NH4HSO4. There are complex relationships governing the formation of these by-products, but they can be minimized by appropriate control of process conditions.
Ammonia Slip
Unreacted NH3 in the flue gas downstream of the SCR reactor is referred to as NH3 slip. It is essential to hold NH3 slip to below 5 ppm, preferably 2–3 ppm, to minimize formation of (NH4)2SO4 and NH4HSO4, which can cause plugging and corrosion of downstream equipment. This is a greater problem with high-sulfur coals, caused by higher SO3 levels resulting from both higher initial SO3 levels due to fuel sulfur content and oxidation of SO2 in the SCR reactor.
Operating Temperature
Catalyst cost constitutes 15–20% of the capital cost of an SCR unit; therefore it is essential to operate at as high a temperature as possible to maximize space velocity and thus minimize catalyst volume. At the same time, it is necessary to minimize the rate of oxidation of SO2 to SO3, which is more temperature sensitive than the SCR reaction. The optimum operating temperature for the SCR process using titanium and vanadium oxide catalysts is about 650–750° F. Most installations use an economizer bypass to provide flue gas to the reactors at the desired temperature during periods when flue gas temperatures are low, such as low load operation.
Catalysts
SCR catalysts are made of a ceramic material that is a mixture of carrier (titanium oxide) and active components (oxides of vanadium and, in some cases, tungsten). The two leading shapes of SCR catalyst used today are honeycomb and plate. The honeycomb form usually is an extruded ceramic with the catalyst either incorporated throughout the structure (homogeneous) or coated on the substrate. In the plate geometry, the support material is generally coated with catalyst. When processing flue gas containing dust, the reactors are typically vertical, with downflow of flue gas. The catalyst is typically arranged in a series of two to four beds, or layers. For better catalyst utilization, it is common to use three or four layers, with provisions for an additional layer, which is not initially installed.
As the catalyst activity declines, additional catalyst is installed in the available spaces in the reactor. As deactivation continues, the catalyst is replaced on a rotating basis, one layer at a time, starting with the top. This strategy results in maximum catalyst utilization. The catalyst is subjected to periodic soot blowing to remove deposits, using steam as the cleaning agent.
Chemistry:
The chemistry of the SCR process is given by the following:4NO+4NH3+O2→4N2+6H2O2NO2+4NH3+O2→3N2+6H2O
The side reactions are given by:SO2+½O2→SO32NH3+SO3+H2O→(NH4)2SO4NH3+SO3+H2O→NH4HSO4Process Description
As shown in FIG. 20, dirty flue gas 112 leaves the power generation system 110. This flue gas may be treated by other air pollution control (APC) subsystems 122 prior to entering the selective catalytic reduction (SCR) subsystem 2170. The flue gas may also be treated by other APC subsystems (not shown) after leaving the SCR and prior to exiting the stack 117. NOx in the inlet flue gas is measured with one or more analyzers 2003. The flue gas with NOx 2008 is passed through the ammonia (NH3) injection grid 2050. Ammonia 2061 is mixed with dilution air 2081 by an ammonia/dilution air mixer 2070. The mixture 2071 is dosed into the flue gas by the injection grid 2050. A dilution air blower 2080 supplies ambient air 152 to the mixer 2070, and an ammonia storage and supply subsystem 2060 supplies the ammonia to the mixer 2070. The NOx laden flue gas, ammonia and dilution air 2055 pass into the SCR reactor 2002 and over the SCR catalyst. The SCR catalyst promotes the reduction of NOx with ammonia to nitrogen and water. NOx “free” flue gas 2008 leaves the SCR reactor 2002 and exits the plant via potentially other APC subsystems (not shown) and the stack 117.
There are additional NOx analyzers 2004 on the NOx “free” flue gas stream 2008 exiting the SCR reactor 2002 or in the stack 117. The measured NOx outlet value 2111 is combined with the measured NOx inlet value 2112 to calculate a NOx removal efficiency 2110. NOx removal efficiency is defined as the percentage of inlet NOx removed from the flue gas.
The calculated NOx removal efficiency 2022 is input to the regulatory control system that resets the ammonia flow rate setpoint 2021A to the ammonia/dilution air mixer 2070 and ultimately, the ammonia injection grid 2050.
SCR Process Controls
A conventional SCR control system relies on the cascaded control system shown in FIG. 20. The inner PID controller loop 2010 is used for controlling the ammonia flow 2014 into the mixer 2070. The outer PID controller loop 2020 is used for controlling NOx emissions. The operator is responsible for entering the NOx emission removal efficiency target 2031 into the outer loop 2020. As shown in FIG. 21, a selector 2030 may be used to place an upper constraint 2032 on the target 2031 entered by the operator. In addition, a feedforward signal 2221 for load (not shown in FIG. 21) is often used so that the controller can adequately handle load transitions. For such implementations, a load sensor 2009 produces a measured load 2809 of the power generation system 110. This measured load 2809 is sent to a controller 2220 which produces the signal 2221. Signal 2221 is combined with the ammonia flow setpoint 2021A to form an adjusted ammonia flow setpoint 2021B, which is sent to PID controller 2010. PID 2010 combines setpoint 2021B with a measured ammonia flow 2012 to form an ammonia flow VP 2011 which controls the amount of ammonia supplied to mixer 2070.
The advantages of this controller are that:    1. Standard Controller: It is a simple standard controller design that is used to enforce requirements specified by the SCR manufacturer and catalyst vendor.    2. DCS-Based Controller: The structure is relatively simple, it can be implemented in the unit's DCS and it is the least-expensive control option that will enforce equipment and catalyst operating requirements.SCR Operating Challenges:
A number of operating parameters affect SCR operation:    Inlet NOx load,    Local molar ratio of NOx:ammonia,    Flue gas temperature, and    Catalyst quality, availability, and activity.
The operational challenges associated with the control scheme of FIG. 20 include the following:    1. Ammonia Slip Measurement: Maintaining ammonia slip below a specified constraint is critical to operation of the SCR. However, there is often no calculation or on-line measurement of ammonia slip. Even if an ammonia slip measurement is available, it is often not included directly in the control loop. Thus, one of the most critical variables for operation of an SCR is not measured.            The operating objective for the SCR is to attain the desired level of NOx removal with minimal ammonia “slip”. Ammonia “slip” is defined as the amount of unreacted ammonia in the NOx “free” flue gas stream. While there is little economic cost associated with the actual quantity of ammonia in the ammonia slip, there are significant negative impacts of ammonia slip:                    Ammonia can react with SO3 in the flue gas to form a salt, which deposited on the heat-transfer surfaces of the air preheater. Not only does this salt reduce the heat-transfer across the air preheater it also attracts ash that further reduces the heat-transfer. At a certain point, the heat-transfer of the air preheater has been reduced to the point where the preheater must be removed from service for cleaning (washing). At a minimum, air preheater washing creates a unit de-rate event.            Ammonia is also absorbed in the catalyst (the catalyst can be considered an ammonia sponge). Abrupt decreases in the flue gas/NOx load can result in abnormally high short-term ammonia slip. This is just a transient condition—outside the scope of the typical control system. While transient in nature, this slipped ammonia still combines with SO3 and the salt deposited on the air preheater—even though short-lived, the dynamic transient can significantly build the salt layer on the air preheater (and promote attraction of fly ash).            Ammonia is also defined as an air pollutant. While ammonia slip is very low, ammonia is very aromatic, so even relatively trace amounts can create an odor problem with the local community.            Ammonia is absorbed onto the fly ash. If the ammonia concentration of the fly ash becomes too great there can be a significant expensive associated with disposal of the fly ash.                            2. NOx Removal Efficiency Setpoint: Without an ammonia slip measurement, the NOx removal efficiency setpoint 2031 is often conservatively set by the operator/engineering staff to maintain the ammonia slip well below the slip constraint. By conservatively selecting a setpoint for NOx, the operator/engineer reduces the overall removal efficiency of the SCR. The conservative setpoint for NOx removal efficiency may guarantee that an ammonia slip constraint is not violated but it also results in an efficiency that is lower than would be possible if the system were operated near the ammonia slip constraint.    3. Temperature Effects on the SCR: With the standard control system, no attempt is evident to control SCR inlet gas temperature. Normally some method of ensuring gas temperature is within acceptable limits is implemented, usually preventing ammonia injection if the temperature is below a minimum limit. No attempt to actually control or optimize temperature is made in most cases. Furthermore, no changes to the NOx setpoint are made based upon temperature nor based upon temperature profile.    4. NOx and Velocity Profile: Boiler operations and ductwork contribute to create non-uniform distribution of NOx across the face of the SCR. For minimal ammonia slip, the NOx:ammonia ratio must be controlled and without uniform mixing, this control must be local to avoid spots of high ammonia slip. Unfortunately, the NOx distribution profile is a function of not just the ductwork, but also boiler operation. So, changes in boiler operation impact the NOx distribution. Standard controllers do not account for the fact that the NOx inlet and velocity profiles to the SCR are seldom uniform or static. This results in over injection of reagent in some portions of the duct cross section in order to ensure adequate reagent in other areas. The result is increased ammonia slip for a given NOx removal efficiency. Again, the operator/engineer staff often responds to mal-distribution by lowering the NOx setpoint. It should be understood that the NOx inlet and outlet analyzers 2003 and 2004 may be a single analyzer or some form of an analysis array. In addition to the average NOx concentration, a plurality of analysis values would provide information about the NOx distribution/profile. To take advantage of the additional NOx distribution information, it would require a plurality of ammonia flow controllers 2010 with some intelligence to dynamically distribute the total ammonia flow among different regions of the injection grid so that the ammonia flow more closely matches the local NOx concentration.    5. Dynamic Control: The standard controller also fails to provide effective dynamic control. That is, when the inlet conditions to the SCR are changing thus requiring modulation of the ammonia injection rate, it is unlikely that the feedback control of NOx reduction efficiency will be able to prevent significant excursions in this process variable. Rapid load transients and process time delays are dynamic events, which can cause significant process excursions.    6. Catalyst Decay: The catalyst decays over time reducing the removal efficiency of the SCR and increasing the ammonia slip. The control system needs to take this degradation into account in order to maximize NOx removal rate.    7. Rolling Average Emissions: Many regulatory emission permits provide for both instantaneous and some form of rolling-average emission limits. To allow for dynamic process excursions, the instantaneous emission limit is higher than rolling average limit; continuous operation at the instantaneous emission limit would result in violation of the rolling-average limit. The rolling-average emission limit is an average of the instantaneous emissions value over some moving, or rolling, time-window. The time-window may be as short at 1-hour or as long a 1-year. Some typical time-windows are 1-hour, 3-hours, 24-hours, 1-month, and 1-year. Automatic control of the rolling averages is not considered in the standard controller. Most NOx emission permits are tied back to the regional 8-hour rolling average ambient air NOx concentration limits.
Operators typically set a desired NOx removal efficiency setpoint 2031 for the SCR and make minor adjustments based on infrequent sample information from the fly ash. There is little effort applied to improving dynamic control of the SCR during load transients or to optimizing operation of the SCR. Selecting the optimal instantaneous, and if possible, rolling-average NOx removal efficiency is also an elusive and changing problem due to business, regulatory/credit, and process issues that are similar to those associated with optimal operation of the WFGD.
Other APC processes exhibit problems associated with:                Controlling/optimizing dynamic operation of the process,        Control of byproduct/co-product quality,        Control of rolling-average emissions, and        Optimization of the APC asset.        
These problems in other processes are similar to that detailed in the above discussions of the WFGD and the SCR.